Conversion of hydrocarbons



AP 3, 1945- R. M. coRN'l-'oRT'H CONVERSION OF HYDROCARBONS Filed June 29. 1940 Y Patented Apr. 32.1945.

.umn-:o sfrArEs PAT ENT.; OFFICE coNvEirsIoN oF maocAnnoNs msnm M. coi-nimh, om, ma.,

Standard Oil Compan ration of Indiana assigner y, Chicago, Ill., a corpo- Application June 29, 1949, serai No, 343,223

(ci 19e-.49)

scams. This invention relates to the production of high4v antiknock gasoline suitable for aviation fuels or as a blending stock'tlierefor, and relates Vmore particularly to the conversion of low octane number hydrocarbon mixtures to high octane number fuels of balanced distribution as regards boiling range -and volatility, suitable for use as aviation gasoline or as a blending stock in the production of aviation .fuels.\

It has been discoveredthat the isomerization `o'f'high-boiling hydrocarbon mixtures is not pracf f tical or desirable, due to the -fact that catalyst life is short and the yield is low, but that the isomerization of low-boiling hydrocarbons takes place4 readily with. good yields. lt has also been discovered that the catalytic reforming of such high-boiling stocks succeeds admirably in converting low anti-knock hydrocarbon mixtures to a `blend of hydrocarbons having greatly increased octane number but that such a process, which is y essentially an aromatization reaction, does not eect a similar beneficial conversion of the lowerboiling hydrocarbons. Y

v It is'a'n object of this invention to provide a process for the conversion of low' octane number liquid hydrocarbons to high octane number liquid hydrocarbons suitable iorruse as premium c fuels. Another object of my invention/is to pro'- Y vide a process for the conversion or liquid hydrocargbons of gasoline boiling range containing a substantialamount of straight-chain hydrocarbons to hydrocarbons oi similar boilingr range rich in branched-chain and/or aromatic hydrocarbons. A further object of my invention is to provide a process for the catalytic conversion 3 comparatively free under hydrogenlpressure of low octane number' hydrocarbons to gasoline of increased octane number.' An additional object of my invention is to provide` a process whereby high octane number hydrocarbon'mixtures suitable dfor blending with other hydrocarbon mixtures of varying octane number properties can be obtained -from low octane number stocks unsuitable for blend'- ing. Further objects and advantages of my proc'- ess will become apparentas .the description.

thereof proceeds.

Briefly described, my process contemplates the conversion of low octane number naphtha to high octane number motor fuels by the catalytic isomerization oi, the lower boiling hydrocarbons to branched-chain paraiiinic hydrocarbons and the t catalytic reforming ofthe higher boiling hydro-E y carbons to aromatic hydrocarbons, both reactions beingcarried out in 'an atmosphere of hydrogen,

the hydrogen formed during the aromatization being sumcient to supply the isomerization reaction, as well 'as the aromatization reaction,

with the necessary hydrogen pressure.

The single drawing which forms a part of this speciiication illustrates in the form of a simplified ilow diagram apparatussuitable for carrying out my process.

Referring no'w to the drawing: Feed stock enters through line Il 'and is directed to fractionator Il by pump I2; As feed stock I canuse, for

example, a petroleum fraction containing virgin gasoline which is substantially propane-free'and which has an end point of about 300 to 400 F. This can be -a light virgin naphtha or ,straightrungasoline or Amay be distilla gasoline or similar hydrocarbon mixtures recovered from distillate we lls or natural gas Wells. Such light virgin naphtha or straight-run gasoline or distillate gasoline or similar hydrocarbon mixn tures, all of whichl are substantially saturated,

are called paraillnic naphtha" as distinguished from' oleiln and aromatic-containing cracked naphtha obtained by catalytic or thermal'conversion processes .but such parailinic'naphthas may, o f course, contain` naphthenes as well as parainns. Such gasolines and naphthas are, generally speaking, oi' low octane number, lbeing characterized by the 'presence ofLconsiderable amounts of straight-chain hydrocarbons and hydrocarbons and aromatics which impart' high octane number,I characteristics to a motor fuel excluded from the high boiling fraction. Topcooling means I3 and bottom-heating means Il can be used tb aid in eilectuating the separation. I can maintain fractionator Il, for. example, at a gage'pressure of from about 0 pounds per square inch to 25 pounds per square inch, preferably '5 pounds per square inch, with a top temperature 0f fr0m 130 F.' to 190 F., preferably A 150 F., and bottom temperature of from 160 F. to 310 preferably 200 F. The lighter hyof highly branched-chain that the heptanes are excluded from the' low boiling traction and that methyl pentanes are opening valve to 175 F. pass overhead through line I5 and, if desired, can be directed through line I6 by I'I therein to debutanizer i8 equipped With top-cooling means I9 and bottomheating means 20 wherein hydrocarbons having four carbon atoms per molecule are taken overhead through line 2l. The butane can be discarded from the system by opening valve 22 in line 23 or can be used for blending with the finished product to provide additional volatility if necessary by opening valve 24 in line 25. 1t can also be converted by means ofvsuch processes as dehydrogenation, polymerization, alkylation, etc., to high octane number gasoline-like products; can be used for blending with butanedeficient motor fuels; or can be similarly utilized in various processes known to those skilled in the art.

The debutanized light naphtha is withdrawn through line 25 and can be directed as such to an isomerization process by opening valve 27 in line 28 or can be depentanized by opening valve 29 in line 30 which leads to depentanizer 3l, Under suitable conditions of temperature and pressure a fractionation can be carried out in depentanizer 3l whereby the major portion of the The heavier products from depentanizer 3l can be withdrawn by opening valve 38 inline 39 but preferably, since these too are of low octane number, they are directed to the isomerization reaction by opening valve 4I inline 40 which joins lino 42. In the event that there was insuliicient isopentane in the light naphtha cut in debutant/.er I8 to Warrant the separation thereof from the remaining hydrocarbons, the debutanized light naphtha from debutanizer I8 will be directed via line 28 to line 112 and thence to an isomerization reaction. Alternately, it is often undesirable or yunnecessary to fractionate further the light hydrocarbons having ve carbon atoms per` molecule pass overhead through line 32 while the higher boiling hydrocarbons are withdrawn through line 33. Depentanizer 3l can be suitably maintained with a top temperature of from about 140 to about 220 F., preferably about 190 F., and a bottom temperature of from about 220 to about 300 F., preferably about 270 F. at a gagepressure of from about to 75 pounds per square inch, preferably 50 pounds per square inch. Top-cooling means 36 and bottom-heating means 35 aid in this fractionation. The pentane -in line 32, which will contain both isopentane and normal pentane can be further fractionated in fractionator 35 to separate the isopentane from the normal pentane. Fractionator 3E can be maintained with a top temperature of about 140 F. to about 200 F., preferably about 175 F., and a bottom temperature of about 155 F. to about 215 F., preferably about 190 F., at a gage pressure of from about 25 pounds per square inch to about 75 pounds per square inch, preferably about pounds per square inch.

Isopentane is a particularly valuable product for use in aviation motor fuel, due both to its high octanenumber and its volatility. Many aviation fuels, particularly those prepared by the oatalytic production of isooctane from butane and butylenes are deficient in low-boiling high octane number hydrocarbons and require the addition of more volatile stocks in order to obtain a balanced fuel.A Normal pentane on the other hand, While of suicient Volatility, does not possessthe desirable high octane number characteristics so naphtha from fraetionator II and accordingly the entire fraction isdirected to line l2 by Opening valve 5l in line 52 which connects line I5 and line i12. By the proper manipulation of valves Il, 5I, 29, 21,41, 33, 49 and 41 it is possible to clirect to the isomerization reaction a light nephtha which may be butane-free, pentane-lree, isopentane-free, or may contain any or all of these constituents.

The light naphtha fraction from fractionator II as previously described passes from line l2 through coils 53 in heater 50 wherein the temperature of the naphtha is elevated to within the range of from about 100 to about 450 F. Hydrogen from any suitable source, and preferably from the aromatization reaction to be described later, enters through line 55 which joins the hydrocarbons in line 56 prior to reactor 5l. The isomerization reaction is carried out under relatively high total pressures, for example, from about 250 to about 3000 pounds per square inch, preferably about 500 to about 1500 pounds per square inch. Of this total pressure, the partial pressure of hydrogen is from about'50 to about 2500 pounds per square inch, preferably about 400 to about 1000 pounds per square inch. The hydrogen need not be a pure product but may contain such impurities as methane, ethano'., etc., in which case the total pressure can be somewhat higher than that specified above.

The heated hydrocarbons pass from coil 53 through line 56, together with hydrogen from line 55, to reactor 5l. Catalyst from line 58 is added to the reactants and hydrogen halide from line 59 can be added to the reactor to promote the reaction. As catalysts can employ an aluminum halide which can suitably be aluminum chloride or aluminum bromide in anhydrous form or can that We prefer to convert it by isomerization to opening valve 45 in line '46 it can be directed for blending with the other products from our process as will be described later.

be the catalyst complex formed during the reaction or from a previous treatment of parainic, naphthenic or even olenic hydrocarbons with aluminum chloride or aluminum bromide in the presence of a hydrogen halide. The aluminum chloride or bromide is preferably introduced into the reaction zone in the form of a slurry or in suspension in, for example, a portion of the feed stock to the reactor. The concentration of the catalyst can vary within rather wide limits depending primarily upon the temperature, reaction time and the catalyst activity. As hydrogen halides I can suitably use hydrogen chloride or hydrogen bromidenor I can employ organic chlorides such as ethyl chloride, propyl chloride orv the corresponding bromides, etc., which under the conditions present will break down to yield hydrogen halides.

'Isomerization reactor 51 can be maintained at the proper temperature by a. jacket 60 about reactor 51 through which flows a heating medium from line 6I, discharging kthrough line 62. A stirrer 63 within the reactor is employed to produce the necessary intimate contact between Y tion reactor f pass by line 9| through cooler Il (if desired) to separator I9 wherein the catalyst settles and separates iromthe hydrocarbons, hy-A drogen and'hydrogen halide, and can be withdrawn through line 91. Ii spent for isomeriza- .tion'the catalyst can bedischarged through line .69 lby opening valve 99'therein. If, however, it is still active for isomerization, the catalyst can be recycled to reactor l'l'by opening valve 1| in line 1| which joins line 58. By the proper manipulation of valves 69 and 10, a part oi' the catalyst can be continuously withdrawn and the remainder recycled to supplement the fresh catalyst entering through line 5B.

The hydrocarbons pass 'from separator 66 lvia line 12to high pressure separator .13 wherein the `hydrogen -andfhydrogen chloride are released and pass overhead through line 'I4 whence they sary only when starting up the process, the hyprocess with dehydrogenation and aromatization of straight-chain hydrocarbons, there is riot only no hydrogen consumption, although the process is carried out Ain an atmosphere of hydrogen, but actually a 4net gain in hydrogen. Accordingly, hydrogen from an extraneous source is ynecesdrogen produced during the reaction being suiiicient to supply the necessary hydrogen atmosphere by recycling. In fact,` since there is a continual increase in hydrogen asthe reaction continues, sumcient excess hydrogen is obtained\to can be recycled to reactor '51 by opening valve 'l1 in line 19 which joins line 55 prior to pump 19 which forces the hydrogen, together with any hydrogen halide, into line 56 at the pressures desired for carrying out the reaction. Generally' speaking, some hydrogen will be consumed in maintaining the catalyst activity, so that all of the exit hydrogen can be recycled, supplemented by hydrogen from an extraneous source, preferably excess hydrogen from the aromatization process. High pressure separator 'I3 can be .maintained at a gagev pressure of ,about 200 pounds per square inch and a temperature of from 60 F, to 120 F., preferably'100 F.

'I'he isomerized hydrocarbons are withdrawn from high pressure separator 13 by line 80 and satisfy the isomerization reaction `hydrogen requirements.

Since the reaction is a dehydrogenation and aromatization reaction rather than av hydrogenation reaction, it is markedly endothermic. vThe charge and hydrogen are generally heated to a temperature above'the desired average temper-- ature in the reaction catalyst chamber to supply the endothermic heat of reaction and this is necessarily true Vif an unheated reactor is used,` the temperature differential depending upon the proper conditions, the apparatus, its size and design. To maintain more uniform temperature and to minimize any purely thermal conversion, the hydrogen can. be heated to a higher temperature than the charge and then can, if desired,

.be injected at multiple points in the catalyst bed.

The catalyst chamber 91 as shown has' catalyst tubes |06 connected with headers |01 and |98,

. the space between the tubes being heated to compensate for the endothermicity of the reaction may be discharged from the system by opening valve 8| in line 82, whereby they can be utilized for blending with high octane number stocks of 1,

insuiiicent volatility or with low octane number stocks to increase the octane number thereof. They can also (and preferably forY our process) be directed to fractionator 83 by opening valve fin line 85 which joins line 86 leading to fractionator 83.

The bottom stock from fractionator having an initialof about 150 to 175 F. and an end point J of about-300 to 400 F., is withdrawn through line 91 and canbe withdrawn from the system for blending 01T. for use as motor fuels by opening valve 89 in line 89. However, since thisfraction is deiicientin high octane number constituents aswell as light ends" it is desirable to convert -these hydrocarbons to hydrocarbons o1'k gasoline Abo range having an increased ocft'ane number. Accordingly the fraction can be directed to a-'catalytic reforming process byopening valve 90 in line 9| which joins line 92 leading to coils 92 in heater 94,'and thence by valve 85 in line 98 to catalytic chamber 91. Simultaneously with the introduction of the hydrocarbon fraction, hydrogen fromany suitable source can be introducedby line 98 and valve99, and by 'opening valve |00 inline |0| is directed through -such as titanium, cerium, thorium and vanadium.

. ably from about 50 to 300 pounds per square inch.

hydrocarbons. The aromatization process can b e carried out at temperaturesbf from about 875 to 1075 F.,

and maintain the required temperature. This can be accomplished by passing hot ii'ue gas or other heating medium in through duct |09 and out through duct ||0. 'As illustrated, a station ary catalyst bed is employed, but this can, with equal suitability, be a moving bed reaction chamber or can employ powdered catalyst. Various designs and modincation will occur readily to those skilled in the art, and I do not intend to be limited t0 this specific design.

As catalysts, I prefer the oxides of the metals v of the left-hand column of group VIof the periodic table, particularly chromium, molybdenum 4and tungsten, but I can also' usegother metallic 'oxides and/or metallic compounds particularly oxides'of the metals of the lefthand column of groups IV and V of the periodic table,

These catalytic oxides can be used alone or on various supports including` magnesia and particularly alumina, especially an activated alumina or an alumina gel. Mixed catalyst can also be used, for example, a mixture of chromium oxide and molybdenum oxide alone or on an lumina support. Another catalyst which can be used is magnesium chromite, either alone or onc a support such ,as the above-mentioned alumina. My catalysts can include any which promote the dehydrogenation and preferably about 890 F., and at gage pressures o f from 30 to 450 pounds per square inch, prefer- Anotherv important operating variable is the time factor" which may be defined' as the amount of time in hours required to put through the catalyst' a volume of feed (measured as liq- Y uid) equal to the volume of the catalystchamber, the volume of the catalyst chamber being'the over-all volume of that portion of the chamber cyclization of Aaliphatic 4 which is filled with the.catalyst. In other words, time factor is the volume of catalyst space divided by the feed rate in volume per hour. I have found that the time factor should be between 0.1l and 25 and preferablybetween 0.2 and 20A K ultimate object and will beaccomplished under the vconditions hereinabove set forth, there will be a net gain of hydrogen by the process, so that complete recycle of all of the hydrogen from the process is neither desirable nor permissible. The process effected under the conditions set forth in this paragraph with a catalyst as set forth in the previous paragraph is called Jhydroforming and the term hydroforming as employed in the appended claims is. hereby defined to mean this process.

The charge and hydrogen from catalyst reactor 91 pass through line and cooler ||2 to high pressure separator' I3 from which the hydrogen (usually containing some light hydrocarbon gases) passes out through line I3. As was previously remarked, there will be a production -of hydrogen during the catalytic reforming so` that not all of it can be recycled to the catalytic reforming process. Accordingly, a part of it can b e discarded by opening valve ||5 in line H3 l while the remainder is recycled by opening valve ||l in line ||8 and valve H3 in line |20 which joins line 98. Valve ||9 should be so adjusted that the mol ratio .of hydrogen to feed stock is maintained within the limits of from 0.5 to 8. Rather than discard the hydrogen through line |65 it is preferable for our process that the hydrogen be directed through valve |2| in line |22 to line 55' which leads to the isomerization reactor. If desired, the hydrogen in line |22 can be purified in hydrogen purifier |23 (shown generally) prior to its use in the isomerization reaction but this is'not ordinarily essential, since the presence of minor amounts of light hydrocarbon gas is not bons of aviation. and motor fuel gasoline boiling range is withdrawn from fractionator 83 by line |33 and can be rerun by opening valve |34 in line' |35 which leads to rerun tower |36. In rerun tower |36 a separation is made between products of a boiling range suitable for aviation gasoline and the heavier products which can be incorporated in motor fuel and which, generally speaking, are too heavy to be' used in aviation gasoline except aviation safety fue The aviation fuel is withdrawn overhead via line |31 and the motor fuel withdrawn via line |38.v On the other hand it may be possible to utilize the combined products as such, and these can be withdrawn by opening valve |39 in line |33. In the event that isopentane was previously separated from the light fraction this can also be blended with the product from fractionator 83 via lines it and |3| which join line |33 and the blended product withdrawn through line |42. Isopentane can also be blended with the aviation fuel in line |3| from rerun tower |33, by diverting it through line |03 and valve IM. By regulating valves and |00 in lines |3| and llllrespectively, a part can be directed to the aviation fuel and a part to the heavier blended fuel. Also, butane from line 25 caribe blended with the products of de-isopentanization, aromatization and isomerization to increase the volatility of the nal product by opening valve 23 in line 25 which joins line |32.

lAs a specific example of my process, I may separate a debutanized light virgin naphtha into two fractions, one having `an end point of about 158 F. and the other anfend point of about 350 F. The lighter fraction is contacted with an aluminum chloride-hydrocarbon complex catalyst prepared by the reaction of aluminum chloride on an aromatic-free light naphtha in the presence of hydrogen chloride. Hydrogen chloride is added 40 to the light hydrocarbon fraction and catalyst,

detrimental to the proper functioning of the isomerization reaction. Any surplus of hydroger-r not required by the isomerization reaction can of course be discarded through line IIS. l

The catalytically reformed heavy naphtha is discharged from high pressure separator ||3 by line |24 andA can be withdrawn for blending with other products or for use as such by opening valve |25 in line |26. However, since this product is more or less deficient in the more volatile gasoline-range hydrocarbons, I prefer to direct it via line |21 through valve |28 to line 86 which leads to fractionator 83. If the isomate from the isomerization reaction has also been directed via line 85 to fractionator 83, the` products of the two processes will be simultaneously fractionated, the butan'es discarded overhead through line |29 while the polymers heavier than aviation gasoline or motor fuel are withdrawn through line |30. Fractionator 83 can be maintained with a top temperature of about tozabout 170 F., preferably about F., and a b" ture Aof about 500 to about 600i@v` preferably about 550? F., and at a gage pressure of about 60 pounds per square inch to about 120k pounds per square inch, preferably about 100 pounds per square in. Top-cooling means |3| and bottomheating means |32 can be used to assist in the fractionation. A fraction comprising hydrocarand the whole thoroughly intermingled by stirring at a temperature of 330 F. for about one hour, at a gage pressure of about 1000 pounds per square inch, a part of said pressure being supplied by hydrogen. Simultaneously the heavy fraction is contacted-with a catalyst comprising molybdenum oxide on alumina at a pressure of 200 pounds per square inch and at a temperature of 980 F. with a time factor of 1.0 hour. ZIhe pressure is supplied in part by hydrogen. The products from both reactions, after the release of fixed gases, arecommingledand fractionated to yield a hydrocarbon mixture of gasoline boiling range of high nantilrnock characteristics.

Although I have described my process as regardsI certain apparatus, for the sake of clarity and simplicity certain details have been omitted; for example, top-cooling coils in any or all of the fractionating towers can be replaced by supplying reiiux to the towers from an outside source or by cooling and condensing the top products 'from the fractionator towers and returning a porimilarly, in place tlf fractionators,

vI can withdraw a portion'o-f thefheavy product,

heat it to increase the temperature sufciently,

. and return the heated products to the fractionom temperaator whereby heat is supplied to .the products to .be fractionated. Also, I have omitted certain details as regards1 pumps, heat exchangers, coolingy means, pressure release valves, etc., all of which,

will occur readily to one skilled in the art an which would naturally be used in -any commercial.

plant employing my process. Moreover, although I have illustrated my process as having only one ,isomerization reactor and one aromatization refraction having an initial boiling point from-about actor, it is contemplated that more than one reactor can be employed,.either in series or in parallel, in either reaction whereby separate regeneration of the catalyst as well as increased capacity can be attained without a "shutdown.

It will be apparent from the above description that I have provided an improved process for the production of aviation gasoline from low octanenumber naphthas,v wherein the lighter constituents are catalytically isomerized to yield branched-chain hydrocarbons, and the heayier constituents which are isomerized only with diif culty or with low yields are catalytically aromatized to yield maximum quantities'of highoctanenumber gasoline-range fuels. Either of these products is suitable for blending with other stocks to increase the octane number and also to supply certain deficiencies in boiling rangeand/or volatility; together, the two products yield a high octane number balanced fuel of good lead response and of suitable boiling range and volatility.-

Moreover', by the combination of these processes, we can eliminate the necessity of an outside hydrogen supply after the system comes to equilibrium, vsince under the conditions set forth, the aromatization process not only supplies suilcient hydrogen for carrying out the dehydrogenationcyclization reaction, but suillcient excess hydrogen is available to supply the isomerization reaction, making the entire process and self-sufficient one.

My process is also ladvantageous in that little or no diluent is charged to the isomerization process, particularly if the isopentane is eliminated from the light hydrocarbons. It has` lbeen found that an equilibrium between the converted isomeric hydrocarbons and unconverted straightchain hydrocarbons is set up during catalytic isomerization, the conditions of operation and the catalyst detennining this toa large extent. However, if isoparaifinic hydrocarbons are charged to the isomerization reactor, the amount of isom-.. erization of straight-chain hydrocarbons is decreased in direct proportion to the amount of isoparamnic hydrocarbons originally present. Consequently, if the isopentane is fractionated from y r thelight naphtha cut, this valuable hydrocarbon can be used in the blending of aviation or motor fuels, while an equivalent amount of normal pentane can be converted to isopentane. l By my process-I can eillciently convert lowerboiling straight-chain hydrocarbons to the more desirable branched-chain hydrocarbons with optimum yield and long catalyst life, and convert vhigher boiling straight-chain hydrocarbons. to high antiknockfaromatics, which, when frac.

tionated simultaneously, yield a. high antiknock gasoline of suitable boilingrange and Volatility.

Although I have described my invention in relationto-certain embodiments thereof, it should be understood that this is by way of illustration .and not a limitation thereon, my invention being limited only as set forth in the appended claims.

I claim:

j 1. A process for the production of high octane number gasoline yfrom a hydrocarbon fraction having a boiling range of from4 about 30 F. to

350 F. and containing a -substantial amount of open-chain paraffinic hydrocarbons, which comprises separating said hydrocarbon fraction into a fraction rich in hydrocarbons having four carbon atoms per molecule,A a fraction rich in isopentane, an isopentane-free fraction having an end point from about 150 to about 175 F. and a \an in/tegrated` contacting said isopentane-free fraction having an end point fromabout 150 to about 175 F.

- with an 'isomerization catalyst in the presence of l hydrogen under conditions adapted to convert a substantial part of the paraftlnic hydrocarbons contained therein to branched-chain paramnic hydrocarbons with consumption of hydrogen. con- .noting soia ifsouon having on inn-.iai boiling point from about 150 to about 175 F. with a catalyst adapted to promote the conversion of the open-chain hydrocarbons contained therein to aromatic hydrocarbons with the production of l5 hydrogen, ata temperature of from about 875- to about 1075 F., a 'pressure oi' from about 3 0 to about 450 pounds per square inch and a time factor of from about 0.1 to about hours in the presence of from about 0.5 to 8-mols of hydrogen 20 per. mol of said fraction having an initial boiling point of about 150. to about 175 F., separating said 'hydrogen from said aromaticV hydrocarbons, 'directing at least asubstantial portion or said hydrogen in amounts necessary to supply the hy'- 25, drogen consumed therein to said contacting step for the conversion of straight-chain hydrocarbonsto branched-chain hydrocarbons, blending said isopentane-rich fraction, said branchedchain paratlinic hydrocarbons and said aromatic oh'ydrocarbons, and increasingthe volatility of said blended hydrocarbons by adding at least a. portion of `said'iraction rich in hydrocarbons having four carbon atoms per molecule.v

`2. A vprocess accordingto claim 1 wherein '55 said conversion oflparailinic hydrocarbons .to

branched-chain hydrocarbons is carried ot in the presence of an aluminum'halide catalyst se'- lected from the grop cnsistingoi' aluminum chloride, aluminum bromide, an aluminum chloride-hydrocarbon complex formed by the reaction voi! anhydrous aluminum chloride with a hydrocarbon in the presence of a hydrogen halide and drocarbons which is substantially free from oleilns and which boils within the naphtha boiling range, `which process comprises obtaining from said charging stock a iirst fraction consisting essentially oi' butanes, asecond-fraction rich in isopentane, a third iraction'higher boiling than isopentane and having an end point in the gen-- eral vicinity of about 150 to 175 F. and a fourth fraction higher boiling than the third fraction and having an initial boiling point within lthe f general vicinity of 150 to 175 F., contacting said third fraction with an isomerization catalyst under conditions .for eiecting substantial isomerization for the conversion of parainic hydrocarbons contained therein to parailinic hydrocar- 'bons oi more highly branched-chain structure, contacting said fourth fraction with an aromatizing catalyst to produce aromatization products 7o proximate range of`30 to 450 pounds per square inch with a time factor of from about .1 to about,

an aluminum bromide-hydrocarbon complex at a temperature within the'approximate range of 875 to 1075 at-a pressure within theapv .fourth fraction,-separating hydrogen from the aromatizationproducts and recycling at least a substantial portion of said hydrogen to said aromatizing catalyst contacting step, recovering a fraction rich in aromatics from the aromatization products, recovering a highly branched-chain.

paraflinic fraction from the products of the isomerization catalyst contacting step and blending at least a portion of each of said last-named two fractions with at least a portion of said fraction a hydrocarbon charging stock which contains a substantial amount of open-chain paraiiinic hydrocarbons, which is substantially free from olens and which boils Within the naphtha boiling range, which process comprises fractionating said charging stock to obtain a first low boiling fraction, a second intermediate boiling fraction having an end point in the general vicinity of about 150 to about 175 F. and a third high Aboiling fraction having an initial boiling point above 150 F., hydroforming said third fraction to produce higher octane number hydrocarbons along with a net production of hydrogen, isomerizing the second fraction with a halide isomerization catalyst under conditions for effecting substantial conversion of paranic hydrocarbons contained therein to paramnic hydrocarbons of more higltLly branched-chain structure, eiecting said isomerization in the presence `of added hydrogen at superatmospheric pressure whereby isomerization catalyst activity is maintained with la net consumption of hydrogen, recovering a fraction rich in aromatics from the hydroforming step, recovering a lowerboiling branched-chain paraflin asvavri fraction from the products of the isomerization step and blending at least a portion of each of said 'last-named fractions with at least a portion of said low boiling fraction to produce a gasoline of balanced volatility and of -high antiknock rating. y

5. The process of claim 4 which Aincludes the step of purifying at least a part of the hydrogen produced in the hydroforming step and employing hydrogen thus purified insaid isomerization step.

6. In a process'for the production of high antiknock rating gasoline of balanced volatility from a hydrocarbon charging -stock which contains a substantial amount of open-chain parainic hydrocarbons, which is substantially free from olens`4 and which boils within the naphtha boiling range; which process comprises hydroforming a higher boiling fraction'of said charging stock, isomeri'z'ing an intermediate boiling ,fraction of said charging stock and subsequently blending at least'a portion of the products of the hydroforming and isomerization steps respectively with lower boiling hydrocarbons to produce a gasoline of balanced volatility and of highantiknock rating, the improved method of voperation which :comprises charging stock to remove lower boiling hydrocarbons therefrom and to'e'ect a cut point between an intermediate boiling fraction charged to isomerization and a higher boiling fraction charged to hydroforming iin the general vicinity of about 150 to 175vF. whereby both heptanes and heavier hydrocarbons are substantially excluded from the isomerization step While methyl pentanes are eX- cluded from the hydroforming step and converted into more highly branched-chain structure in the isomerization step.

' ROBERT M. CORNFORTH.

initially '.fractionating said 

